It is rare in a chemical process that the process uses stoichiometric amounts of reactants with essentially complete conversion in a simple reactor. Hence, where the reagents constitute a significant cost of the process, the unreacted material is often recycled to the reactor, usually after some physical separation of the desired product from the un-reacted material. Sometimes this separation can be achieved internally within the reactor, for example, where the reactants are gaseous, the product is in liquid form at reaction conditions and is withdrawn continuously, and a stirred tank reactor with gas induction impellers is utilized. In this particular situation, the physical separation and the recycle occur within the reactor vessel.
Alternatively, the separation and recycle of reactant can take place external to the reactor. One example of this configuration would be a plug flow gas phase reactor where the product can be condensed from the gas phase by cooling. The unreacted gases may then be re-compressed, and at least partially returned to the inlet of the reactor, perhaps after other conditioning, such as purification or chemical separation.
There are several reasons why the amounts of reactants utilized in the reactor to form the desired end product are rarely stoichiometric. It could be that vapor pressure limitations require a non-stoichiometric reaction. For example, in high-pressure, gas-phase hydrogenations of a high boiling organic the hydrogen will be present in large excess even though a high single pass conversion of the reactant is theoretically possible. An alternate reason would be that the reaction is equilibrium limited. For example in acid catalysed esterifications the alcohol is frequently in excess to achieve high conversion of the acid.
While it may be possible to achieve high conversion of the reactants in an equilibrium limited reaction to the desired end product economically utilizing a large excess of reactants, an alternate possibility is the removal of one of the reaction products. For example, gas stripping could remove the water from an esterification reaction to continually move the conditions in the reactor out of or away from equilibrium and thereby drive the reaction toward full conversion.
However, where a product cannot be removed in-situ to drive the equilibrium toward full conversion, then a high overall conversion is likely to be achieved only by separation of the product from the reactant mixture and subsequent recycle of the un-reacted material to the reactor.
Additionally, even where an equilibrium reaction has certain conditions or aspects that are favorable for high conversion to the desired end product, the kinetics of the reaction may suggest that a higher overall production rate or better process economics can be achieved by running a reactor at conditions favoring a relatively low conversion of the reactants and then recycling the un-reacted material after physical separation of the product. Some exemplary reactions or processes of this type where conversion of the reactants to the desired product is only partial, and in which significant quantities of un-reacted material remain that can be recycled to the reactor after physical separation of the product include the reaction of synthesis gas to methanol, di-methyl ether, mixtures thereof, Fischer-Tropsch waxes, and ammonia.
Using the equilibrium limited reaction for the production of methanol as an example, because methanol is one of the largest by volume chemicals produced in the world today, the conversion to methanol is typically carried out in a two-step process. In a first step, methane is reformed with water or partially oxidized with oxygen to produce carbon monoxide and hydrogen, with some carbon dioxide and residual methane, (i.e., synthesis gas or syn-gas). In a second step, the syn-gas is converted into methanol.
The second step of converting the syn-gas into methanol is a well-known process. It typically involves a catalytic process using a copper-based catalyst, such as a catalyst comprising a reduced zinc oxide/copper oxide mixture, among others. To provide the optimum production of methanol from this reaction, the reaction is typically carried out at pressures within the range of 40-100 bars and at temperatures in excess of 200 degrees C. and below 320 degrees C., with a temperature range of between 220 and 280 degrees C. being most common. The production of the syn-gas itself is typically carried out at pressures within the range of 20-40 bars depending on the reformer technology that is utilized.
Due to the particular mechanism of the reaction for the production of methanol, the reaction does not go to completion, as the concentration of produced methanol is limited by equilibrium. Specifically, the amount of methanol contained in the product gas exiting the reactor comprises about 6-8 mol % of the total gas, although it can be higher. This methanol is removed from this product gas stream by condensing it through cooling the product gas stream to below 110 degrees C., and most commonly below 60 degrees C. The cooled methanol can then be removed from the gas stream while the excess syn-gas is sent back to the reactor in order to further react the excess syn-gas. This enables additional methanol to be obtained from the syn-gas recycled back to the reactor in combination with an amount of fresh syn-gas that is also charged to the reactor.
In performing this recycle step, one well-known process involves the use of a recycle compressor which receives the excess syn-gas from the separator and compresses it in order to overcome the pressure drop that occurs within the reactor and separator. This type of reactor is commonly referred to as a recycle loop reactor and is schematically shown in FIG. 1. In this reactor, the concentration of methanol in the syn-gas leaving the reactor is low enough that the volumetric flow rate of the excess syn-gas through the recycle compressor is typically two (2) to ten (10) times the volumetric flow rate of the fresh syn-gas being introduced into the reactor separately from the syn-gas charged to the reactor from the recycle compressor. A purge gas stream of approximately 4 to 8 percent of the recycled syn-gas stream is also expelled from the recycle loop, prior to re-compression to control the concentration of inert material that builds up in the reactor as a result of the recycle.
One significant drawback to the recycle loop reactor described above is the cost of the recycle compressor. Often the recycle function is incorporated into a single drive train compressor that compresses the syn-gas up to the pressure of the recycle loop and also provides for re-compression of the recycle gases. The compression train is an expensive item of equipment and may be the single most expensive purchased component in the construction of a methanol producing facility. As stated previously, a high recycle flow is utilized to enable high overall conversion of syn-gas to be achieved. The recycle compressor also becomes significantly cheaper per unit volume compressed as the scale of the plant is increased. Thus, the use of a process using a recycle compressor is, as a practical matter, preferential for those facilities producing a relatively high daily output of methanol relative to the current maximum single train production plant, such as those facilities producing in the neighborhood of 5,000 tons per day in 2005, where maximum efficiency is required and integration of the syn-gas compressor and recycle compressor can be achieved in order to make the use of a recycle compressor economically viable.
As an alternative to a recycle loop reactor, it has been conceived that the facility could utilize a multiple reactor set or cascade process (FIG. 2) whereby the syn-gas is initially fed to a first reactor for reaction with the catalyst contained therein to produce methanol. The product gas then enters a first separator or knock-out pot wherein the methanol produced in the first reactor is cooled into liquid form and separated from the excess syn-gas. The first separator can operate to separate the methanol for the syn-gas in any desired manner, such as by gravity or by applying centrifugal focus to the products. The remaining excess syn-gas is then fed to a second reactor, which undergoes the same reaction, thereby producing additional methanol. The additional methanol is removed from the second reactor and directed to a second separator in the same manner. The number of reactors and separators can be selected to create a multiple reactor set that achieves the desired conversion percentage of the syn-gas to methanol. For example, when using an optimal syn-gas composition that has a conversion to methanol of 50 percent in each reactor, the reactor set can be selected to include four reactors and four separators, which theoretically results in the achievement of a 95 percent overall conversion of the syn-gas to methanol after the fourth separator.
However, in more realistic situations, when the syn-gas composition is less optimal, such as when the stoichiometry of the syn-gas is away from the stoichiometry of the reaction desired, e.g., the ratio ([moles hydrogen]−[moles carbon dioxide])/([moles carbon monoxide]+[moles carbon dioxide]) is between 2.5 and 3.0, then more than four reactor sets and perhaps as many as 10 reactor sets are required for high conversion (>95%). Additionally, when there are stoichiometric quantities of reactant in the syn-gas, but high levels of inerts also present, then a high number of reactor sets will be required, such as where autothermal reforming with air is performed for the production of syn-gas resulting in dilution of the syn-gas with high levels of nitrogen.
There are several areas in a methanol production process where the ability to employ a cascade of reactors would be considered beneficial. Principally the benefit from using the cascade comes from not requiring a recycle compressor. Furthermore, it should be borne in mind that in a grass roots methanol production plant the recycle compressor is often part of the compressor associated with syn-gas compression. Thus, eliminating the syn-gas compressor along with the recycle compressor in a grass roots installation gives the maximum benefit.
While the use of a simple cascade of reactors for this particular purpose is disclosed in certain prior art references, the references that discuss the use of such a cascade of reactors focus exclusively on methods by which the reformer operating pressure can be matched to the methanol synthesis pressure. For example, U.S. Pat. Nos. 5,177,114; 5,245,110; 5,472,986; and 7,019,039 are each patents that disclose inventions in the field of autothermal reforming using air rather than oxygen. However, while these patents very generally disclose the use of cascade reactors in the methanol production process, they do not address the issues of how the cascade of reactors can be made cost effectively. Furthermore, each of U.S. Pat. Nos. 5,177,114; 5,245,110; and 5,472,986 disclose a methanol production process where the recycle compressor can be eliminated as a result of operating a reformer autothermally, and then converting the syn-gas to a methoxy compound using three to five reactor sets with product condensation between each stage. Recognizing conventional wisdom that a cascade process cannot achieve a high syn-gas to methanol conversion, carbon efficiencies for the methanol synthesis section of less than 80% are quoted, whereas in conventional plant efficiencies in excess of 95% are achievable.
Additionally, U.S. Pat. No. 6,255,357 discloses a methanol production process that uses pressurization of the oxidant gas for fired heating of the steam reformer as a means of achieving a mechanically feasible high pressure steam reformer with an operating pressure sufficient to ensure a sufficient operating pressure throughout the process. The process also includes a cascade of reactors downstream from the reformer in which the reformed syn-gas is converted to methanol. The pressurization of the incoming natural gas into the reformer avoids the requirement of a syn-gas compressor upstream of the cascade of reactors, as well as avoiding the need for a recycle compressor. However, as with the previous references, the cascade of reactors is only very generally disclosed without any discussion as to how the cascade can be made economically.
Other situations regarding a methanol production process where it would be considered advantageous to avoid use of a recycle compressor include those where the compressor would be added to an existing methanol production facility in a retro-fit capacity, or as part of an addition to a planned methanol production facility construction for the purpose of removing any remaining methanol from the purge gas discharged from the facility. One system of this type that addresses the loss of the potential and actual methanol present in the purge gas stream is disclosed in U.S. Pat. No. 6,258,860, which is incorporated by reference herein in its entirety. The process disclosed directs the purge gas stream produced by a methanol synthesis zone to another methanol synthesis or production zone in order to both collect the methanol present in the purge gas stream as well as to further react the unreacted components of the purge gas stream to produce additional methanol.
However, the process disclosed in the '860 patent has certain drawbacks in that it utilizes a compressor to compress the combined purge gas and recycled syn-gas stream prior to further reacting the combined stream. Because, as discussed previously, the recycle compressor is the highest cost item in a methanol production system, the use of additional recycle compressors to recover methanol from a purge gas is highly undesirable, especially for systems producing a relatively low daily output of methanol relative to the current maximum single train production plant, such as those facilities producing in the neighborhood of 5,000 tons per day in 2005.
Another example where additional compressor capacity may be avoided through the use of a set of cascade reactors could occur as part of a re-vamp or de-bottle-necking of a methanol plant. If the re-vamp or de-bottlenecking entails increased syn-gas production, then the capacity of the methanol converter would be required to be increased. It may be possible to increase the effectiveness of the reactor through better packing of catalyst or dividing the catalyst into multiple beds in a single reactor. However, where the reactor already makes effective use of the catalyst, it may not be possible to economically increase the performance of the reactor. Further, another limitation on the operation of the reactor in this situation is the pressure drop of the process gas across the reactor. Increasing the re-circulation rate, increasing the catalyst volume, increasing the feed flow or reducing the purge rate will all increase the pressure drop of the process gas through the reactor. There will of course be a consequent limitation on the capacity of the re-circulation compressor as well to recompress the gas for reintroduction into the reactor.
One alternative to mitigate these issues would be to operate the methanol converter at reduced conversion conditions, but with a higher gas feed rate, thereby allowing pressure drop limitations to be avoided and then utilizing a separate cascade reactor system to convert the un-reacted gases to methanol without the requirement of an additional compressor or replacing the original reactor. This also has the advantage of being a lower risk method of increased throughput, as the original reactor performance is well known. There will of course be many other circumstances under which a cascade system can be utilized, but all of these circumstances will be reliant on a cost-effective design of cascade system.
One significant drawback with the multiple or cascade reactor set types utilized in methanol production as described in the prior art results from the end use construction of each reactor, heat exchanger(s) and separator forming the individual reactor set. Specifically, because a portion of the syn-gas is lost in each reactor set based on its conversion to methanol, often each subsequent reactor set and separator is constructed to be smaller than the immediately preceding ones to accommodate the reduction in the flow rate of incoming syn-gas. This might initially be anticipated to be highly beneficial based upon the reduction in the amount of material necessary to construct each successive reactor set. However, each reactor set requires the same functionality, connections, cooling and access for catalyst replacement, which become more difficult and/or expensive to manufacture on a progressively smaller scale. In addition, the cooling, gas-liquid separation and re-heating of the methanol-bearing stream as it passes between the various reactor sets must be effected in an energy efficient and cost effective manner. Further, all of the reactor sets and separators must be constructed to be operable at the elevated pressures (40-100 bars) that the reactions occurring for the conversion of the syn-gas to methanol require.
One example of a system that attempts to address this shortcoming is disclosed in U.S. Pat. No. 6,723,886 in a methanol production process using reactive distillation. However, while there is removal of methanol between reactor beds by condensation within the reactor, the condensation takes necessarily takes place at reaction temperature, and condensation at elevated temperature limits the conversion of methanol to approximately 60%. However, even with the significant restriction this places on methanol production, this is in accord with the current industry view that condensation at reduced temperature is not viable.
Therefore, it is desirable to develop a multiple or cascade reactor set and a process for the production of products of equilibrium limited reactions, e.g., methanol, using the multiple reactor set to obtain a high percentage conversion of feed syn-gas to methanol by condensing the methanol in the reactor effluent in an interstage feed/effluent heat exchanger. It is also desirable that the multiple reactor set be operable without the need for a gas recycle compressor and preferably without the need for the construction of multiple individual reactors, heat exchangers, and separators. In other terms, the heat exchanger design should be suitable for efficient operation and integration into the reactor sets, while also minimizing the number of necessary equipment items.
With regard to the goal of minimizing the necessary number of equipment items in a reactor set, it is easier to understand the conventional approach to solving this problem of eliminating equipment items and reducing the cost of equipment items by reference to the specific problems of a conventional methanol synthesis loop. Apart from the recycle compressor, a methanol synthesis loop contains six principle operations: 1) pre-heat of the gas; 2) reaction of the gas to form methanol; 3) removal of the heat of reaction as high grade heat; 4) cooling of the gas to methanol condensation temperatures; 5) condensation of the methanol using cooling water; and 6) vapor/liquid separation. In a typical plant there may be two integrations of these functions for the purposes of minimizing the necessary equipment items which are removal of the heat of reaction is performed by steam raising in a shell and tube reactor, and pre-heat of the gas by feed-effluent exchange. Thus, a typical synthesis loop will consist of at least six equipment items: 1) a start up heater; 2) a feed/effluent heat exchanger; 3) a reactor; 4) a high grade heat recovery unit; 5) a water cooler; and 6) a gas-liquid separator.
Steam raising directly in the reactor does eliminate the requirement for a separate high-grade heat recovery unit. However, it also requires a steam drum with the reactor and so does not reduce the number of equipment items.
With regard to the use of the feed/effluent heat exchanger, the highest energy efficiency is achieved with a high effectiveness heat exchanger that is able to maximize the cooling of the effluent stream. Increasing the amount of high grade heat recovered reduces the temperature difference in the feed/effluent exchanger. Therefore, for maximum high grade heat recovery a high effectiveness heat exchanger is required. However, shell and tube heat exchangers as used in prior art multiple reactor sets can only achieve high effectiveness through the coupling of multiple heat exchanger units, again increasing the number of equipment items required. The usefulness of high grade heat recovery is, in part dependent, on the temperature at which it is recovered. In particular, for a methanol process the high grade heat recovery from the methanol synthesis section is utilized for steam raising for the reformer. This requires that the stream from which heat is being recovered is above a minimum temperature, typically 200-250 deg C. However methanol condensation temperatures are in the region of 60-100 deg C. For efficient operation, therefore, heat exchangers are required that can operate with a hot gas temperature span of approximately 150 deg C. The heat of reaction is recovered by cooling the reactant stream by the equivalent of typically 50-100 deg C. of sensible heating. If the feed gas is introduced to the methanol reactor at a temperature below high grade heat recovery temperature this represents a loss of energy efficiency in the system and increases the low grade cooling requirement. Consequently the temperature difference in the feed-effluent exchanger will be kept to less than 50 deg C. and typically 20-30 deg C. Where high single pass conversions can be achieved in the reactor, such as with a balanced stoichiometry, high operating pressure, efficient heat removal or a low overall conversion the temperature constraints may be more relaxed. However, this often brings greater reactor complexity or lower overall efficiency.
The performance measure of a heat exchanger can be described in terms of temperature span and log mean temperature difference between stream. The value (span divided my lmtd) is referred to as NTU count and as can be seen above it would be desired for an energy efficient methanol process that the fee/effluent exchangers would operate with an NTU count above 5, and more preferably above 7.
The problem concerning the number of equipment items is also not alleviated when a recycle loop is replaced with a cascade system. With no recycle there is no requirement for a recycle compressor. However, for each contact with the catalyst there will be up to six additional equipment items, as discussed previously. One option to reduce the number of equipment items is to eliminate some of the heat exchangers. For example, instead of recovering high grade heat from the reactor gases, the gases can be used to directly heat the incoming feed gases. The feed/effluent exchanger is then smaller as a result of an increased driving temperature, but the reaction heat is then lost to the cooling water and a less efficient process is produced.
Therefore, to improve the economics and efficiency of the prior art methanol cascade systems, it is necessary to solve the following issues: 1) to minimize the number of equipment items; 2) to increase the effectiveness of the feed/effluent exchangers; and 3) to integrate multiple functions into single equipment items.